Hydrocarbon conversion process



April 26, 1966 SPILLANE ETAL 3,248,437

HYDROCARBON CONVERSION PROCESS Filed Jan. 3. 1961 2 Sheets-Sheet 1 FIGURE 1 Z 1. 4g llllll INVENTORS LEO J. SPILLANE y ROY L. GRANTOM April 26, 1966 Filed Jan. 3. 1961 L. J. SPILLANE ET AL HYDROCARBON CONVERSION PROCESS 2 Sheets-Sheet 2 FIGURE 2 3o 25 '30 45 I 30 ll J 32 4 3| 39 INVENTORS LEO J. SPILLANE BY ROY L.GRANTOM ATTORNEY carbons.

United States Patent 3,248,437 HYDROCARBON CONVERSION PROCESS Leo J. Spillane and Roy L. Grantom, El Dorado, Ark., assignors to Monsanto Company, a corporation of Dela ware Filed Jan. 3, 1961, Ser. No. 80,243 7 Claims. (Cl. 260666) The present invention relates to a process for the noncatalytic dehydrogenation of cyclic hydrocarbons. More particularly, the present invention relates to a process for the non-catalytic partial oxidation of cyclo-parafiin hydrocarbons to cyclo-olefin hydrocarbons.

Both catalytic, and non-catalytic methods of causing the dehydrogenation of cyclo-paraflin hydrocarbons to aromatic hydrocarbons are known to the art. The reaction is a stepwise dehydrogenation which, if stopped at the appropriate time before going to completion, will produce cyclo-olefins, e.g., monounsaturated and/ or diunsaturated. Unfortunately, however, it is rather difficult to immediately stop the reaction at any desired degree of un saturation. Even if reaction conditions may be controlled to a degree that cyclo-parafiin hydrocarbons are not dehydrogenated all the way to aromatic hydrocarbons while in the reaction chamber, the further problem of cooling the products with sufficient rapidity to prevent further reaction arises. Thus, presently, the dehydrogenation of cyclo-paraffin hydrocarbons cannot normally be stopped short of the production of aromatic hydrocarbons when the reaction is carried out at the most generally effective temperatures, whether by catalytic or non-catalytic means.

It is an object of the present invention to provide a process for the non-catalytic dehydrogenation of cyclic hydro- A further object of this invention is to provide a process whereby cyclo-paraflin hydrocarbons may be partially dehydrogenated to a coresponding cyclo-mono-olefin hydrocarbon. Another object of this invention is to provide a process for the dehydrogenation of cyclo-parafiin hydrocarbons whereby the degree of dehydrogenation may be effectively controlled. A specific object of the present invention is to provide a process for the non-catalytic partial dehydrogenation of cyclo-parafiin hydrocarbons to cyclo-rnono-olefin hydrocarbons wherein significantly less cyclo-diolefin or aromatic hydrocarbons are produced. Additional objects will become apparent from the description of the invention herein disclosed.

In fulfillment of the above objects it has been found that substantial amounts of cyclic mono-olefins may be produced in relatively good yields by dehydrogenation through the non-catalytic partial oxidation of cyclo-parafiin hydrocarbons in a reaction unit such as that shown in FIGURE 1 and under the conditions hereinafter set forth.

The apparatus and operation of this invention is more apply described and illustrated by reference to the accompanying drawings. FIGURE 1 presents a cross-section of the reaction unit which is preferred in the practice of this invention. FIGURE 2 presents a flow diagram of the entire process including ,the reaction unit of FIG- URE 1.

Referring first to FIGURE 1, a preheated liquid hydrocarbon feed stream is passed into the mixing head 2 by means of line 1. Preheated oxygen or air is brought into the mixing zone through line 3. The mixed hydrocarbons and oxygen are then forced into the reaction chamber 5 through a spray nozzle 4 which increases uniformity of mixing of the feed fuels entering into the reaction chamber. The spray nozzle has the added advantage of preventing flashbacks, premature ignition, etc. The hydrocarbon and oxygen mixture sprayed into the reaction chamber is ignited by an igniting surface 6. The igniter may be a spark plug or any other suitable device ice and may be placed anywhere in the reaction chamber but must remain so located that the incoming oxygen and hydrocarbon will strike its surface. Once combustion is initiated the igniting surface is of no further use since the reaction is autogenous. It is not necessary that an igniting surface be used since spontaneous ignition will occur if the oxygen and feedstock are heated to a sufiiciently high temperature prior to entry into the reactor.

In FIGURE 1 is shown an alternate arrangement for the introduction of the oxygen or air into the reaction chamber. This alternate arrangement allows the heat of the reaction chamber to be used for preheating purposes. This is accomplished by introducing the oxygen into the reaction chamber through line 7 and annular passage 8. If this alternate arrangement is not used, the annular passages are closed off at the elbow 9 and thus they become insulating aids. 4

The reaction chamber may be lined with any suitable refractory material. It also may be of any shape and size consistent with good flow characteristics and certain limitations as to length to width ratio discussed later in the specification. The reaction chamber of this process is characterized by having a constricted opening at its exit end, the optimum amount of constriction varying with feedstocks and other considerations.

The reaction product of the combustion of the hydrocarbon feed in the reaction chamber next passes through the constrictedopening 11. Generally, for flow and other design characteristics, it is desirable to use a converging and diverging nozzle arrangement as exemplified by the sloping lines leading to the constricted opening. It is to be understood, of course, that the degree of slope and the ratios of slope of convergence to slope of divergence are not to be in any way limited by the drawing in FIGURE 1. On passing through the constricted opening the gases are expanded into the exhaust chamber 12. The rapid expansion of the gases in the exhaust chamber brings about a proportionally rapid lowering of the temperature of the gaseous reaction products. The divergence of the exhaust chamber may be so designed that the expansion of the hot product gases will suficiently quench the efiluent gaseous reaction product stream thus alleviating any need for further cooling. However, a coolant such as water may be injected into the exhaust stream to aid the expansion quench. A coolant has the added advantage of slowing the high velocity gaseous products and simultaneously removing the heat generated by the slowing action. This coolant may be injected through line 13 or may be injected through the walls at an angle to the path of the exhaust stream or may be injected co-current with the exhaust stream. The liquid products which are condensed from the exhaust stream are collected in tank 15 and are removed by line 16. The lighter non-condensed gases exit the exhaust chamber through line 14.

FIGURE 2 presents the flow arrangement of the proc-- ess of this invention. The hydrocarbon mixture storage or source is represented by 21. flows by line 22 into a preheater 2 3 and then by line 24 into the mixing zone 25. The oxygen or air is taken from storage orsource 26 by means of line 27 into preheater 28 and then by line 2 9 into the mixing zone 25. If a diluent --gas such as steam, nitrogen, etc., is used, it may be introduced into the oxygen line through line 30. Also, if it is desired to use the alternate arrangement of oxygen entry described in the explanation of FIGURE 1, the oxygen stream may 'be fed into the reaction chamber by means of line 50. The reactants pass from the mixing zone 25 through the spray nozzle 30 into the reaction chamber 31, where they are brought into contact with igniting surface 32 if such is used. These products are passed through constricted opening 33 and expanded into the exhaust chamber 34. In this chamber Feed from storage cooling of the hot products takes place, either by expansion of the hot gases or by both expansion and injection of a coolant or by injection of a coolant alone. The liquid products collect in tank 35 and are removed by line 36. Non-condensed product material flows from the exhaust chamber by line 67 into a secondary condensing unit 38 where the remaining liquid product is condensed from the product stream. From this second condenser, the gaseous product flows to its future utility by means of line 39 and the liquid products are removed and collected by means of line 40. The showing of only one condenser in the drawing is not limiting since the number of condensers is entirely dependent upon the efficiency of the expansion quench and the secondary condenser and the overall degree of condensation desired. The liquid product from condenser may be maintained separately or may be passed by line 41 into admixture with the liquid products being removed by line 36 from collection tank 35. The combined mixture then may be passed into line 43 and thereby into distillation unit 44 or it may be passed direct to storage by line 49. The liquid product passed into the distillation unit 44 is separated into a light, intermediate, and heavy overhead fraction and a still heavier bottoms fraction. These four fractions exit the distillation unit by lines 45, 46, 47, and 48 respectively.

It will become readily apparent to those skilled in the art that a great number of variations and modifications of the equipment and flow arrangement as presented in FIG- URES 1 and 2 may be made without departing in any way from the spirit and scope of this invention. It is to be understood that within the preceding process description the use of the word oxygen is to include air also.

To further illustrate the invention herein disclosed, the following examples are presented. It is to be understood, of course, that these examples are in no way to be construed as limiting the application, operation, or conditions of this invention.

Example I The reaction unit used in this example was constructed of stainless steel with a fused alumina lining on the internal surface of the reaction chamber. This reaction unit was similar to the one presented in FIGURE 1. The length of the reaction chamber was /2 inches and the internal diameter 1 (one) inch. At the exit end of the reaction chamber the walls of the chamber converge at an angle of 30 to a constricted opening of 0.1 inch diameter. The angle of divergence from the constricted opening was approximately 7.5 for a distance of 0.4 inch. At this point the diverging slope ceased and the constriction opened into .an exhaust chamber having a diameter of 8 inches and a length of 24 inches.

A hydrocarbon feedstream comprised of cyclohexane containing about 5 mol percent of cyclic and paraffinic hydrocarbon impurities was introduced into the above described reaction unit. The operating conditions for this reaction are as follows:

Hydrocarbon flow rate, liquid liters/minute 0.1 Air flow rate, gaseous liters/minute 40 Molar ratio of oxygen to hydrocarbon feed in reaction chamber 0.38 Temperature of hydrocarbon at entry into mixing head, C 370-395 Temperature of oxygen at entry into mixing head, C. 575-585 Reaction temperature, C. 550-600 Reaction pressure, p.s.i.g. i 58-585 Run length, minutes 20 Residence time, sec. 110x10 Flow velocity, ft./ sec 3.79

The product of the reaction chamber was then passed through the venturi constriction and expanded in the exhaust chamber, thus Condensing the liquids from the product stream. 'No secondary condenser was used. The 1 4.- liquid product was recovered and represented 64.5 weight percent of hydrocarbon charge recovered. The liquid product was then distilled into 4 fractions, a 35.2 to 80.2 C. fraction representing 5.8 volume percent of the liquid product, an 80.2 to 829 C. fraction representing 65.1 volume percent of the liquid product, and 82.9 to 99 C. fraction representing 5.5 volume percent of the liquid product and a bottoms fraction boiling above 99 C. and representing the remainder of the liquid product of the reactor. The 80.2 to 82.9 C. fraction was analyzed and found to contain 14.2 volume percent cyclohexene, 0.3 volume percent benzene and a trace of 1,3 cyclohexadiene with the remainder being cyclohexane. The cyclohexene produced represented 9.1 volume percent of the liquid product.

Example II The reaction unit used in this example was the same as that used in Example I.

A hydrocarbon feedstream comprised of methylcylopentane containing about 5 mol percent of hexane, 2,2- dimethylpentane and 2,4-dimethylpentane impurities was introduced into the above described reaction unit. The operating conditions for this reaction were as follows:

Hydrocarbon flow rate, liquid liters/minute 0.06 Air flow rate, gaseous liters/ minute 28 Molar ratio of oxygen to the hydrocarbon feed in reaction chamber 0.44 Temperature of hydrocarbon at entry into mixing head, C 425-435 Temperature of oxygen at entry into mixing head, C. 525-575 Reaction temperature, C. 520-580 Reaction pressure, p.s.i.g 35-365 Run length, minutes 20 Residence time, sec. 116x10- Flow velocity, ft./ sec 3.95

The product of the reaction chamber was then passed through the venturi constriction and expanded in the exhaust chamber, thus condensing the liquids from the product stream. No secondary condenser was used. The liquid product was recovered and represented 65.8 weight percent recovery. The liquid product was then separated into 4 fractions, a methylcyclopentane fraction representing 76 volume percent of the liquid product, an olefinic fraction representing 12 volume percent of the liquid product, an aromatic fraction representing 4 volume percent of the liquid product, and an oxygenated fraction representing 8 volume percent of the liquid product of the reactor. The olefinic fraction was found to be a mixture of l-methylcyclopentene, 3-methylcyclopentene, and 4-methylcyclopentene with only trace quantities of methylcyclopentane and aromatics.

The feedstock of the present invention is preferably a hydrocarbon mixture containing a substantial concentration of cyclo-parafiin hydrocarbons. The amount of cyclo-parafiins present in the feedstock may range, however, from 1 to 100% weight. Since yields of product are relative to the amount available, cyclo-parafiin hydrocarbons having cyclo-paraflin hydrocarbon concentrations of 50% weight or greater would be more preferred. The molecular weight ranges of the feedstock may vary considerably, primarily being limited by the molecular weight ranges in which cyclo-paraffins occur. However, the preferred molecular weight range would include cycloparafiin hydrocarbons of 4 to 8 carbon atoms in the cyclic nucleus of the molecule. The cyclo-parafiin hydrocarbons which may be dehydrogenated to cyclo-monoolefin hydrocarbon includes alkyl substituted cycloparaffins such as methylcyclopentane, methylcyclohexane, ethylcyclohexane, methylcyclooctane, etc. Further, cycloparaffin hydrocarbons having polyalkyl substituents as 1,3-dimethylcyclohexane, etc., may be effectively dehydrogenated to a corresponding cyclo-mono-olefin hydrocarbon. In the preferred practice of this invention the alkyl substituent to the cyclic nucleus will have 3 or less carbon atoms per alkyl group. Non-limiting examples of non-substituted cyclo-parafiin hydrocarbons which may be converted by the present invention to cyclo-monoolefin hydrocarbons are cyclobutane, cyclopentane, cyclohexane, cycloheptane, cyclooctane, etc. The impurities which may be in the feedstock include normal and isoparaifin hydrocarbons and small quantities of olefin and aromatic hydrocarbons. However, in order that the separation of the product fraction may be simplified, it is preferable that there be no aromatic hydrocarbons present in the feedstock.

The temperatures at which the present process is to be beneficially operated may range from 450 to 700 C. A somewhat more preferred range of reaction temperatures is that of 500 to 650 C. A still more preferred range for operating the present invention will be found between the temperatures of 550 and 625 C. The optimum temperatures will vary somewhat with feedstock composition, molecular weight and the other operating variables but will generally remain within these ranges.

Temperatures in the reaction chamber may be controlled by controlling preheat temperature, feed rates, feed to oxygen ratios, and residence time in the reaction chamber. Diluent gases also aid in reactor temperature control. In general, the temperature of the reactor lags behind the preheated streams until a reactor temperature of 325 to 375 C. is reached. At this point, if no ignitor is used spontaneous ignition occurs. After reaching this 325 to 375 C. temperature the reactor temperature surges rapidly, almost instantaneously to -a new temperature of 450 to 475 C. All other variables remaining fixed, the rate of increase of reactor temperature then decreases somewhat and proceeds to follow roughly the rate of increase of preheat temperature. The effect of feed rates and ratios and residence time in the reactor chamber will be discussed at some length later in the specification.

In producing cyclo-rnono-olefin hydrocarbons with the present invention, reaction chamber pressures of to 100 p.s.i.a. may be utilized. It is more preferred, however, to operate the present invention with pressures within the range of approximately to 70 p.s.i.a.

Another variable having considerable affect on the efficiency of the process herein disclosed is the hydrocarbon feedstock-oxygen ratio. At any given temperature and reaction chamber pressure the yield of dehydrogenated products is greatly dependent upon the oxygen concentration in the reactor. Generally, throughout this application reference is made to oxygen. However, it is to be understood, that this does not limit the practice of the present invention to pure oxygen. On the contrary, it is more practical and much less expensive to use air. Of course, when air is used, it is necessary to take into con sideration the inert components of'the air in determining flow rates, concentration, heat requirements, etc. Though it has been determined that some dehydrogenation will take place in the present process in the complete absence of oxygen, this does not occur to any extent. In its broad application the oxygen concentration may range anywhere within the broad limits of 1 to 80 mol percent. However, a more practical and preferred range of oxygen concentration is 10 to 60 mol percent with a still more preferred range being from 15 to 50 mol percent oxygen. These molar concentrations are based on total feed to the reaction chamber.

Input rates of both the hydrocarbon mixture and the combustion supporting gas have considerable effect on the reactor temperature. As input flows increase there is an increase in temperature. This, of course, may be compensated for to some degree by the manipulation of the other operating variables. The feed flow rates which are within the practical applicabilty of the present invention, as set forth herein, are 0.2 to 10 volumes of feed per minute per volume of reaction space within the reaction chamber. More preferred flow rates for the hydrocarbon feed are found in the range 0.5 to 2.0 volumes of hydrocarbon feed per minute per volume of reaction space. The input flow rates of the combustion supporting gas and diluents will, of course, be dependent upon the desired ratio of these gases to the feed mixture.

The residence time of the reactants in the reaction chamber is limited by temperature, pressure, feedstock, feed flow rate, reactor flow velocities and the reactor design. The last of these, reactor design, is generally fixed and is not, therefore, considered a variable affecting residence time. Residence time may generally vary from 10 10 seconds to 250x10 seconds. A somewhat more preferred range of residence time is 50x10 seconds to 200 10- seconds.

The flow velocity of the reactants through the reaction chamber should be within the broad range of from 1 to 15 feet per second. Velocities of 2 to 10 feet per second will generally yield somewhat better results, however. Within these ranges flow velocities will vary in accordance with the considerations expressed in regard to residence time, e.g., feed rates, temperatures, pressures, feedstocks, and reactor design.

The reaction unit utilized in the process herein described will generally have a cylindrical internal surface. It may be constructed of any material which will withstand the physical requirements of temperature and pressure and will not adversely effect the reaction taking place within.

. Materials adversely affecting the reaction may be used with an internal lining of a suitable refractory material such as fused alumina, hard carbon or graphite. A nonlimiting example of a suitable reaction unit is one constructed of a stainless steel with an internal fused alumina lining.

Inthe design of the reaction chamber, it is of some importance that ,a proper ratio between the length and diameter of the reaction chamber be maintained. This is especially true in processing heavier feedstocks in that if the reaction chamber is overly long in relation to its diameter, a significant amount of carbon and tar formation takes place as a result of secondary condensations in the reaction chamber. The criticality of the length to diameter ratio becomes less as the feedstock becomes lighter or lower boiling. The ratio of length to diameter may vary from 0.5:1 :to 10:1, however, it is somewhat more desirable to use ratios of 1:1 to 6:1. I

The relationship which must exist between the diameters of the constricted exit orifice of the reaction chamber and the reaction chamber is a critical factor in the design of the reaction unit. This relationship has considerable effect on reactor flow velocities, residence time and pressures in that the pressure ratio between the reaction chamber and the exhaust or expansion chamber is primarily controlled by the size of the exit orifice. For this reason the ratio of the diameters of the exit orifice and the reaction chamber is generally expressed as a pressure ratio of the reaction chamber side of the orifice of the exhaust or expansion chamber side. The pressure ratios at which the present process is operable may range from 1.2:1 to 10:1. A more preferred pressure ratio, however, would be within the range of 2:1 to 7:1.

The exit orifice of the reaction chamber generally consists of the orifice itself and a converging and diverging section as shown in FIGURE 1. The angles of convergence and divergence are given as the angle between the converging and diverging slope and the plane of the walls of the reaction chamber. The angle of convergence from the reaction chamber wall to the exit orifice may range from 10 to 90 with a generally more useful and practical range being from 30 to The angle of divergence from the exit orifice wall to the exhaust chamber wall may vary from 5 to a more preferred range being 30 to 75.

The exhaust or expansion chamber of this reaction unit is not strictly defined. It generally, is considered to include the diverging area immediately following the exit orifice. Actually, it is within this diverging section of the exit orifice that expansion begins and in many cases all expansion takes place in this area. The exhaust or expansion chamber then functions as an area where the gaseous products exiting the reaction chamber are rapidly expanded and thereby cooled. The gaseous products on passing through the exit orifice attain high velocities and in expansion in the diverging area of the expansion chamber reach even higher velocities. It is a second function of the exhaust or expansion chamber to decelerate these gaseous products back to speeds practical for condensation and collection of the vaporized liquid products and for control of the product gases.

The size of the exhaust chamber is not critical within broad limitations. It must be large enough in diameter to allow some expansion and of a length sufficient for deceleration of the high velocity gaseous products. Water or other coolants may be injected into the hot high velocity gases to aid both in cooling and in slowing these gaseous products. Injection of coolants may be made countercurrent, co-current, or at an angle to the direction of flow of the gaseous products. The use of coolants will generally allow smaller exhaust chambers in the reaction units. The upper limit in size of the exhaust chamber is governed only by the practical physical size of the chamber.

The exhaust chamber as shown in the drawing, FIG- URE 1, provides a collection tank for condensed liquids at its end. While this is a very useful arrangement for collection of the product condensed within the exhaust chamber it is not intended to be limiting on the present invention in any way. Any number of methods for collecting the liquid products condensed in the exhaust chamber will be readily apparent to anyone skilled in the art.

For introduction of the reactants into the reaction chamber, a spray type mixing head as illustrated by FIG- URE 1 is preferred. However, any of the many variations of this type of mixing head or any other mixing head may be used in the practice of the present invention, though not necessarily with equivalent results. The reactants may also be introduced separately as by the introduction of one of the reactants through annular passages surrounding the reaction chamber as shown in FIGURE 1 of the accompanying drawings. This method of introducing reactants into the reaction chamber has the advantage of utilizing heat produced in the reaction chamber for the preheating of the reactant.

What is claimed is:

1. A process for the non-catalytic dehydrogenation of cyclo-paraffin hydrocarbons to cyclo-mono-olefin hydrocarbons comprising introducing the cyclo-paraflin feedstocks concurrently with oxygen in a mol ratio of 0.1:1 to 2:1 into a reaction unit comprised of 2 chambers connected by a constricted passageway, the first chamber being a reaction chamber and constructed in such manner that the length to diameter ratio of the chamber is 0.521 to 10:1 and the second chamber being an expansion chamber, with the hydrocarbon feed input rate to the reaction chamber at 0.2 to 10.0 volumes of feed per minute per volume of reaction space within the reaction chamber, the temperature within the reaction chamber being in a range of about 400 to 700 C., the reactor chamber pressure being in a range of about 20 to 100 p.s.i.a. and the pressure ratio between the reaction chamber and the expansion chamber being maintained within a range of approximately 1.221 to 10:1, and with the chamber flow velocity being at approximately 1 to 15 feet per second and the residence time of the reactants in the reaction chamber being in a range of about 10 10- seconds to 200x10 seconds, thereafter passing the high temperature, high velocity gaseous products through the constricted passageway into the expansion chamber, thereby effecting rapid cooling through the adiabatic expansion of the gaseous products exiting the constricted passageway and thus causing an instantaneous termination of the reaction, and thereafter decelerating the gaseous products to facilitate collection of the cyclo-mono-olefin products.

2. The process of claim 1 wherein the cyclo-paratfin hydrocarbon feedstock is a C to C cyclo-parafiin.

3. The process of claim 1 wherein the cyclo-paraffin hydrocarbon feedstock is one containing cyclo-paraflin of 5 and 6 carbon atoms in the cyclic nucleus.

4. The process of claim 1 wherein the mol ratio of oxygen to hydrocarbon feed is in the range of 0.2:l.l to 1:1.

5. The process of claim 1 wherein the pressure in the reaction chamber is 40 to p.s.i.a.

6. The process of claim 1 wherein th pressure ratio between the reaction chamber and the expansion chamber is maintained at 2:1 to 7: 1.

7. The process of claim 1 wherein the oxygen is introduced as air.

References Cited by the Examiner UNITED STATES PATENTS 2,692,292 10/1954 Robinson 260666 2,870,231 1/1959 Hughes et al 260683 XR 2,905,731 9/1959 Seed 260683 XR 2,908,733 10/1959 Sage 260679 3,049,574 8/1962 Johnson 260666 DELBERT E. GANTZ, Primary Examiner.

ALPHONSO D. SULLIVAN, Examiner. 

1. A PROCESS FOR THE NON-CATALYTIC DEHYDROGENATION OF CYCLO-PARAFFIN HYDROCARBONS TO CYCLO-MONO-OLEFIN HYDROCARBONS COMPRISING INTRODUCING THE CYCLO-PARAFFIN FEEDSTOCKS CONCURRENTLY WITH OXYGEN IN A MOL RATIO OF 0.1:1 TO 2:1 INTO A REACTION UNIT COMPRISED OF 2 CHAMBER CONNECTED BY A CONSTRICTED PASSAGEWAY, THE FIRST CHAMBER BEING A REACTION CHAMBER AND CONSTRUCTED IN SUCH MANNER THAT THE LENGTH TO DIAMETER RAIO OF THE CHAMBER IS 0.5:1 TO 10:1 AND THE SECOND CHAMBER BEING AN EXPANSION CHAMBER, WITH THE HYDROCARBON FEED INPUT RATE TO THE REACTION CHAMBER AT 0.2 TO 10.0 VOLUMES OF FEED PER MINUTE PER VOLUME OF REACTION SPACE WITHIN THE REACTION CHAMBER, THE TEMPERATURE WITHIN THE REACTION CHAMBER BEING IN A RANGE OF ABOUT 400 TO 700*C., THE REACTOR CHAMBER PRESSURE BEING IN A RANGE OF ABOUT 20 TO 100 P.S.I.A. AND THE PRESSURE RATIO BETWEEN THE REACTION CHAMBER AND THE EXPANSION CHAMBER BEING MAINTAINED WITHIN A RANGE OF APPROXIMATELY 1.2:1 TO 10:1, AND THE CHAMBER FLOW VELOCITY BEING AT APPROXIMATELY 1 TO 15 FEET PER SECOND AND THE RESIDENCE TIME OF THE REACTANTS IN THE REACTION CHAMBER BEING IN A RANGE OF ABUT 10X10**3-SECONDS TO 200X10**3 SECONDS, THEREAFTER PASSING THE HIGH TEMPERATURE, HIGH VELOCITY GASEOUS PRODUCTS THROUGH THE CONSTRICTED PASSAGEWAY INTO THE EXPANSION CHAMBER, THEREBY EFFECTING RAPID COOLING THROUGH THE ADIABATIC EXPANSION OF THE GASEOUS PRODUCTS EXITING THE CONSTRICTED PASSAGEWAY AND THE CAUSING AN INSLANTANEOUS TERMINATION OF THE REACTION, AND THEREAFTER DECELERATING THE GASEOUS PRODUCTS TO FACILITATE COLLECTION OF THE CYCLO-MONO-OLEFIN PRODUCTS. 